Catalytic process for converting renewable resources into paraffins for use as diesel blending stocks

ABSTRACT

A process for converting renewable resources such as vegetable oil and animal fat into paraffins in a single step which comprises contacting a feed which is a renewable resources with hydrogen and a catalyst which comprises molybdenum, a non-precious metal and an oxide to produce a hydrocarbon product having a ratio of even-numbered hydrocarbons to odd-numbered hydrocarbons of at least 2:1.

FIELD OF THE INVENTION

The present invention relates to a process and a catalyst for theproduction of linear and branched paraffins (hydrocarbons) fromrenewable resources, that are useful as a blending stock for diesel fuelfor use in warm and cold climates.

BACKGROUND OF THE INVENTION

The high cost and increased environmental footprint of fossil fuels andlimited petroleum reserves in the world have increased the interest inrenewable fuel sources. Renewable resources include ethanol from cornand sugar for use in automobiles, and plant oils for use as diesel fuel.Research in the diesel fuel area includes two main areas, bio-diesel andgreen diesel.

Transesterification of fatty acids in triglycerides into methyl estersusing methanol and a catalyst such as sodium methylate, produces FAME(Fatty Acid Methyl Ester), which is commonly referred to as bio-diesel.These methyl esters, mainly linear C₁₄ to C₂₂ carboxylic acids, can beused as fuel or can be blended into diesel refined from crude oilsources. The transesterification reaction is complex. To be used asdiesel fuel, costly modification of diesel engine is necessary as wellas conversion of associated piping and injector configurations. Otherdisadvantages include poor performance of bio-diesel in cold weatherapplications, limiting its world wide use to warmer climates, and pooremissions. In addition, use of bio-diesel increases maintenance costsdue to poor lubricity, increased viscosity, and high oxygen content.Bio-diesel, while a renewable resource, brings a high cost of use forprocessing and use in engines.

Diesel from renewable resources, commonly referred to as green diesel,involves converting the fatty acids in triglycerides into linear alkanesvia hydrodeoxygenation (HDO). The triglyceride backbone is converted topropane and separated. Green diesel can be used as a fuel by itself oras a mixture with diesel from petroleum feedstocks (petro diesel) withlittle to no engine modification and can be processed in refineriescurrently refining crude oils. Current processes involve multiple stepsto obtain green diesel fuel with comparable properties with petrodiesel. Steps include hydrodeoxygenation, hydroisomerization and/orhydrocracking.

Delmon, B. “Catalysts in Petroleum Refining 1989” in: Studies in SurfaceScience and Catalysis, Eds. Trimm, D. L., Akashah, S., Absi-Halabi, M.,and Bishara, A. (Elsevier, Amsterdam, 1990), pp 1-38, discloses thetransformation of a very large portion of crude oil to usable productsdepends on cracking and hydrotreating processes. Over the last severaldecades, hydrotreating processes have become more complex anddiversified and includes such processes as, hydropurification (e.g.,removal of sulfur, nitrogen, oxygen, metals, etc.), hydroconversion(e.g., production of jet fuels or lubricants), and hydrocracking (mildor heavy hydrocracking). Specifically, the removal of sulfur, nitrogen,oxygen, and metals are called hydrodesulfurization,hydrodenitrogenation, hydrodeoxygenation, and hydrodemetallization,respectively.

Certain hydrotreating catalysts for use with petroleum feedstockscomprise one or more non-precious metals such as nickel, cobalt,molybdenum and tungsten supported on mono- or mixed-metal oxides such asalumina, silica or silica-alumina. The catalysts can be promoted byGroup I metals (e.g., lithium, sodium and potassium) and/or fluorine,boron, and phosphorus. The catalyst is activated by simultaneousreduction and sulfidation in place before subjecting it to hydrotreatingreactions. Catalysts consisting of molybdenum supported on alpha-aluminawith promoters such as cobalt (Co—Mo/Al₂O₃) or nickel (Ni—Mo/Al₂O₃) areextensively used in the hydrotreating of petroleum fractions and resids.

The catalyst most commonly used for the production of diesel fromrenewable resources comprises a precious metal such as platinum and/orpalladium. Murzin et al. in Industrial Engineering Chemical Research,Vol. 45 (2006) pp. 5708-5715, disclose numerous metals used for suchcatalysis. Platinum and palladium gave the best conversion of desiredproducts. Nickel catalysts produced unwanted heavier products such asdimers due to the recombination of moieties resulting from extensivecracking of the feed material.

Processes are known to produce green diesel. Such processes suffer fromone or more deficiencies such as requiring multiple steps, multiplereactors, different catalysts per step and expensive precious metalcatalysts.

While there have been great efforts in the green diesel research area,there remains a need for a process of hydrotreating of a renewable feedsource where the hydrodeoxygenation, hydroisomerization andhydrocracking processes are simplified using a less expensive,non-precious metal catalyst.

SUMMARY OF THE INVENTION

Provided is a process for hydrodeoxygenation of a renewable resourcewhich process comprises (a) providing a feed which is a renewableresource; (b) contacting the feed with a catalyst in the presence ofhydrogen at a temperature of 250 to 425° C. and a pressure of 500 to2500 psig (3,450 to 17,250 kPa), wherein the catalyst comprises anoxide, molybdenum one or more active metals selected from the groupconsisting of nickel, cobalt, or mixtures thereof, and the catalyst isin sulfided form, to produce a hydrocarbon product having a ratio ofodd-numbered hydrocarbons to even-numbered hydrocarbons of at least 2:1.

The active metals used in the process of this invention are non-preciousmetals including nickel and cobalt, which are used in combination withmolybdenum. Preferably, the catalyst comprises nickel. More preferably,the catalyst comprises nickel and cobalt.

DETAILED DESCRIPTION OF THE INVENTION

All tradenames used herein are shown in capital letters.

The present invention relates to the production of diesel fuel fromrenewable resources wherein hydrodeoxygenation occurs underhydrotreating conditions over a catalyst comprising a non-precious metaland an oxide, wherein the oxide is a support. The process producessubstantially linear saturated hydrocarbons with mild branched and mildcracked moieties. These hydrocarbons can be used as diesel fuel or as anadditive to blend with petro diesel and possess similar characteristicsand performance of petro diesel. Examples of appropriate feed includebut are not limited to vegetable oils, animal oils, and oils from wood.

Renewable Resources

The feed is a liquid feed, which is a renewable resource and can be anyplant or animal derived oils, fats, free fatty acids and the like. Therenewable resources can be any oil, e.g., such as those containingtriglycerides or free fatty acids, where the major component comprisesaliphatic hydrocarbon chains having C₁₂ to C₂₀ moieties. Preferably thefeed is an oil derived from plants and/or animals, comprising one ormore triglycerides. The feed may comprise a mixture of triglycerides.These triglycerides can be derived from a plant selected from the groupconsisting of pine, rape seed, sunflower, jathropa, seashore mallow andcombinations of two or more thereof. The feed can also be a vegetableoil selected from the group consisting of canola oil, palm oil, coconutoil, palm kernel oil, sunflower oil, soybean oil, crude tall oil, andcombinations of two or more thereof. The feed source may also comprisepoultry fat, yellow grease, tallow, used vegetable oils, or oils frompyrolysis of biomass. The feed may also be oils derived from marinelife, such as algae. The process is very flexible and selection of thefeed is based on availability and cost.

Catalyst

The catalyst comprises an active metal, a first oxide and optionally asecond oxide. The active metal is one or more non-precious metals. Thefirst oxide comprises a mono- or mixed metal oxide, zeolite, orcombinations of two or more thereof, and is used as a support. When asecond oxide is used it comprises a zeolite.

In one embodiment of this invention, the catalyst comprises an activemetal and a first oxide, which may or may not be a zeolite. The catalystmay or may not comprise molybdenum.

The active metals can be nickel (Ni), cobalt (Co), molybdenum (Mo),tungsten (W), or mixtures thereof, e.g., nickel-molybdenum (NiMo),cobalt-molybdenum (CoMo). Preferably, the active metal is Ni. Themetal(s) may be either in the reduced or sulfided (e.g., Ni₉S₈, CO₉S₈,MoS₂) form. When the active metal is nickel, higher amounts, such as atleast 40 wt %, are needed to reduce in the presence of alumina as thefirst oxide.

In a reducing step, the catalyst is treated with hydrogen, preferably atelevated temperatures, such as from 100° C. to 400° C. Typically thecatalyst temperature is increased during hydrogen flow, such as startingat a temperature of about 130° C. and increasing to a temperature of250° C. or 350° C. Such methods are known to those skilled in the art. Aparticular procedure for reducing catalyst is provided below in theExamples.

The catalyst may be sulfided by contacting the prepared catalyst with asulfur-containing compound such as thiols, sulfides, disulfides, H₂S, orcombinations thereof at elevated temperatures. The catalyst can besulfided before it is used or during the hydrotreating process byintroducing a small amount of sulfur-containing compounds, such asthiols, sulfides, disulfides, H₂S, poultry fat, or combinations of twoor more thereof, in the feed. Sulfiding may be important for the longterm activity of the catalyst, depending on reaction conditions and feedcompositions. A detailed sulfiding procedure is described below in theExamples.

Optionally, a metal promoter may be used with the active metal in theprocess of the present invention. Suitable metal promoters include: 1)those elements from Groups 1 and 2 of the periodic table; 2) tin,copper, gold, silver, and combinations thereof; and 3) combinations ofgroup 8 metals of the periodic table in lesser amounts.

The active metal, including whether reduced or sulfided, will beselected based on the desired product.

The first oxide comprises a mono- or mixed metal oxide, zeolite, andcombinations thereof used as a support for the active metal. The firstoxide is used as a support for the active metal. Materials frequentlyused as the first oxide are porous solids with high total surface areas(external and internal) which can provide high concentrations of activesites per unit weight of catalyst. The first oxide may enhance thefunction of the active metal; and supported catalysts are generallypreferred because the metal is used more efficiently.

The first oxide comprises a porous solid oxide with high total surfaceareas (external and internal) which can provide high concentrations ofactive sites per unit weight of catalyst. Preferably the first oxide haspores of a relatively small diameter, that is preferably 50 nm or less.Preferred first oxides have a surface area greater than 20 m²/g, morepreferably, the first oxide has a surface area greater than 75 m²/g,still more preferably the first oxide has a surface area of at least 100m²/g. Generally surface area is less than 300 m²/g.

The first oxide can be any porous solid oxide with high surface areaincluding, but not limited to, oxides such as silica, alumina, titania,titania-alumina, titania-silica, calcium oxide, barium oxide, zirconia,lanthanum oxide, magnesium oxide, kieselguhr, silica-alumina, includingzeolites, and zinc oxide. The first oxide is preferably selected fromthe group consisting of alumina, silica, titania, zirconia, kieselguhr,silica-alumina and combinations thereof. More preferably the first oxideis alumina, silica, kieselguhr, or a combination thereof.

The catalyst may further comprise other materials including carbon, suchas activated charcoal, graphite, and fibril nanotube carbon, as well ascalcium carbonate, calcium silicate and barium sulfate.

When the first oxide is not a zeolite, the catalyst may comprise asecond oxide, which comprises a zeolite. The second oxide can bephysically bonded to or mixed with the active metal supported on thefirst oxide.

The second oxide comprises one or more zeolites. The second oxide isparticularly important for hydroisomerization and hydrocracking in theevent the first oxide is not a zeolite. That is, if the first oxide doesnot comprise a zeolite, then a second oxide, which comprises a zeolite,is advantageously used when hydroisomerization and hydrocracking aredesired reaction processes. Zeolites can be generically described ascomplex aluminosilicates characterized by a three-dimensional frameworkstructure enclosing cavities occupied by ions and water molecules, allof which can move with significant freedom within the zeolite matrix. Incommercially useful zeolites, the water molecules can be removed from orreplaced within the framework without destroying its structure. Zeolitescan be represented by the following formula: M_(2/n)O.Al₂O₃.xSiO₂.yH₂O,wherein M is a cation of valence n, x is greater than or equal to 2, andy is a number determined by the porosity and the hydration state of thezeolite, generally from 0 to 8. In naturally occurring zeolites, M isprincipally represented by Na, Ca, K, Mg and Ba in proportions usuallyreflecting their approximate geochemical abundance. The cations M areloosely bound to the structure and can frequently be completely orpartially replaced with other cations or hydrogen by conventional ionexchange.

The zeolite structure is a corner-linked tetrahedra with Al or Si atomsat centers of tetrahedra and oxygen atoms at corners. Such tetrahedraare combined in a well-defined repeating structure comprising variouscombinations of 4-, 6-, 8-, 10-, and 12-membered rings. The resultingframework is one of regular channels and cages, which impart a usefulpore structure for separation. Pore dimensions are determined by thegeometry of the aluminosilicate tetrahedral forming the zeolite channelsor cages, with nominal openings of 0.26 nm for 6-membered rings, 0.40 nmfor 8-membered rings, 0.55 nm for 10-membered rings and 0.74 nm for12-membered rings (these numbers assume ionic radii for oxygen). Thoseskilled in the art will recognize that zeolites with the largest poresbeing 8-membered rings, 10-membered rings, and 12-membered rings areconsidered small, medium, and large pore zeolites, respectively. Poredimensions are critical to the performance of these materials incatalytic and separation applications, since this characteristicdetermines whether reactant molecules can enter and product molecules(in the catalytic application case) can exit the zeolite framework. Inpractice, it has been observed that very slight decreases in ringdimensions can effectively hinder or block movement of particularreactants or catalysis products within a zeolite structure.

Zeolites are available from various sources. A comprehensive listing ofzeolites vendors is disclosed in “CEH Marketing Research Report:Zeolites” by, 2005, Chemical Economics Handbook-SRI International.

Acid forms of molecular sieve sorbents can be prepared by a variety oftechniques including ammonium exchange followed by calcination or bydirect exchange of alkali ions for protons using mineral acids or ionexchangers (for a discussion of acid sites in zeolites see J. Dwyer,“Zeolite, Structure, Composition and Catalysis” in Chemistry andIndustry, Apr., 2, 1984).

Preferred zeolites are selected from those having medium (10-memberedring) and large (12-membered ring) pore groups. More preferably, thezeolite is selected from the group consisting of MFI (ZSM-5), MEL(ZSM-11), FAU (zeolite Y or USY), MOR (mordenite), and BEA (beta).

Physically mixing the active metal or the active metal and first oxide(supported metal catalyst) with zeolites is sufficient, but one can alsoco-extrude or pelletize the supported metal catalyst and zeolite afterintimately mixing the two, if one desires to do so. A benefit ofphysically mixing is that the relative concentrations of metal/firstoxide and second oxide, thus the catalyst composition, can be changed.Varying the catalyst composition changes the composition of the greendiesel product.

The catalyst can be prepared using any of a variety of ways known in theart. Preferably, a preformed (e.g., already calcined) first oxide isused. For example, the first oxide is preferably calcined beforeapplication of the active metal. The method of placing the active metalon the first oxide is not critical. Several methods are known in theart. Many suitable catalysts are available commercially.

Relative proportions of active metal and first and optional secondoxides, while not critical, are important in that if too little activemetal is present, initial activity will be lower than desired and a longactivation period may be required for the catalyst to reach maximumactivity. It will be appreciated that the higher the weight percent ofactive metal, the faster the reaction. A preferred content range of theactive metal in the catalyst is from about 0.1 wt % to about 90 percentby weight of the total supported catalyst. A more preferred active metalcontent range is from about 0.2 wt % to about 75 wt %. A furtherpreferred active metal content range is from about 0.5 wt % to about 60wt %.

In one embodiment of this invention, a hydrocarbon is produced having ahigher ratio of odd-numbered to even-numbered hydrocarbons. In thisprocess, the active metal preferably comprises nickel, more preferably,does not comprise molybdenum. The nickel content of the catalyst forthis process is at least 40 wt % (combined nickel and nickel oxide).Preferably, the nickel content is 40 wt % to 90 wt %, more preferably,45 wt % to 60 wt %.

In one embodiment of this invention, a hydrocarbon is produced having ahigher ratio of even-numbered to odd-numbered hydrocarbons. In thisprocess, the active metal preferably comprises nickel, cobalt andmolybdenum. The nickel content of the catalyst for this process isgenerally between 0.2 wt % and 20 wt %, more preferably, between 0.5 wt% and 15wt %.

The zeolite can be present in any amount. Preferably the zeolite ispresent in an amount of at least 10 wt %, based on the total catalystweight, to achieve some hydrocracking and hydroisomerization.Preferably, the zeolite is present in an amount of at least 25 wt %,more preferably 25-50 wt %.

Hydrotreating Process

The hydrotreating process may comprise, but is not limited to, threemajor reactions: hydrodeoxygenation (HDO), hydroisomerization (HI)and/or hydrocracking (HC). It is known in the art that minor reactionscan occur during these steps without significantly altering the desiredproduct.

The HDO process is the removal of oxygen from the fatty acids intriglycerides and other free fatty acids to produce a paraffin(hydrocarbon) product. The HDO can occur either as a decarbonylation,decarboxylation or hydrodeoxygenation or a combination thereof.Decarboxylation refers to the process of removal of oxygen as carbondioxide producing a paraffinic hydrocarbon. Decarbonylation refers tothe process of removal of the oxygen as carbon monoxide and waterdirectly creating an unsaturated hydrocarbon or indirectly by addinghydrogen to produce a saturated hydrocarbon. Hydrodeoxygenation refersto the process of removal of oxygen as water by adding hydrogen. Indecarboxylation and decarbonylation, the resulting paraffinichydrocarbon is one carbon unit shorter than the corresponding carboxylicacid. In hydrodeoxygenation, the resulting hydrocarbon has the samenumber of carbons as the corresponding carboxylic acid.

Advantageously, the present invention may be tailored to control theroute of oxygen removal. For processes that desire minimal use ofhydrogen, the decarboxylation and direct decarbonylation routes can beused. For a process that desires minimal evolution of carbon monoxideand carbon dioxide, the indirect decarbonylation or hydrodeoxygenationare the preferred routes.

Controlling the route of oxygen removal also impacts the chain length ofthe paraffins in the treated feed. The chain length may play animportant role in determining which particular deoxygenation process touse. For renewable resources having 18-carbon chains, there may be apreference for n-heptadecane (product of decarbonylation ordecarboxylation) or n-octadecane (product of hydrodeoxygenation withhydrogen consumption). n-Heptadecane (C₁₇) has a lower melting pointthan n-octadecane (C₁₈), which in turn affects the cold-performancecharacteristics of the diesel blending stock. Additionally, producingC₁₇ removes oxygen from the fatty acid primarily as CO and/or CO₂(reduced hydrogen consumption) whereas making C₁₈ hydrocarbons removesoxygen primarily in the form of H₂O (reduced greenhouse gas emissions).Depending on the conditions, one may prefer C₁₇ or C₁₈ hydrocarbons.These routes can be selectively controlled by varying the type and/orcomposition of the catalyst as described herein.

Straight chain hydrocarbons, specifically in C₁₇ to C₁₈ chain lengths,provide good cetane numbers but possess poor cold weather capabilities.Hydroisomerization and hydrocracking improve the cold weatherproperties. In hydroisomerization, a straight chain hydrocarbon isconverted into a branched hydrocarbon. Preferably, isomerization iscontrolled so that the branched hydrocarbon or the mixture of linear andbranched hydrocarbons boils in the range of petro diesel. Hydrocrackingreduces the chain length. Shorter hydrocarbons provide a lower meltingcomponent in green diesel or as an additive to petro diesel. HI and HCboth substantially improve the cold weather properties of green dieselby lowering the cloud and pour points.

Combining process steps, as in one alternative described herein, isadvantageous for numerous reasons. Successfully combining the stepsdecreases the need for additional investment. A single reactor is usedin the present invention to produce the desired diesel rather than a twostep process in separate reactor vessels. Surprisingly, in the presentinvention, a single catalyst can be used in contrast to a conventionaltwo step process, which uses two different catalyst compositions, onefor the HDO and one for HI/HC. Furthermore, the steps of pretreatment offeed material and removal of impurities from the intermediate product,after HDO and prior to HI/HC, can be eliminated.

Hydrotreatment, as described in the present invention, comprisescontacting the feed with hydrogen at elevated temperatures and pressuresin the presence of the disclosed catalyst compositions, tohydrodeoxygenate, hydroisomerize and/or hydrocrack the feed into thedesired fuel. Temperatures range from 250 to 425° C., preferably at 300to 400° C., most preferably from 325 to 375° C. Pressures range from 500to 2500 psig (3,450 to 17,250 kPa), preferably 1000 to 2000 psig (6,900to 13,900 kPa).

In an alternative embodiment there is provided a process forhydrodeoxygenation of a renewable resource which process comprises (a)providing a feed which is a renewable resource; (b) contacting the feedwith a catalyst in the presence of hydrogen at a temperature of 250 to425° C. and a pressure of 500 to 2500 psig (3,450 to 17,250 kPa),wherein the catalyst comprises one or more active metals selected fromthe group consisting of nickel, cobalt, molybdenum, tungsten andcombinations of two or more thereof and a first oxide, wherein thecatalyst is in reduced form, to produce a hydrocarbon product having aratio of odd-numbered hydrocarbons to even-numbered hydrocarbons of atleast 2:1, preferably at least 3:1, more preferably at least 5:1 andmost preferably at least 10:1. When the renewable resource comprisesover 50% C₁₈-based components, such as triglycerides, the processgenerally favors decarbonylation and/or decarboxylation rather thanhydrodeoxygenation.

In a second alternative embodiment there is provided a process forhydrodeoxygenation of a renewable resource which process comprises (a)providing a feed which is a renewable resource; (b) contacting the feedwith a catalyst in the presence of hydrogen at a temperature of 250 to425° C. and a pressure of 500 to 2500 psig (3,450 to 17,250 kPa),wherein the catalyst comprises molybdenum and one or more active metalsselected from the group consisting of nickel, cobalt, or mixturesthereof and the catalyst is sulfided prior to use, to produce ahydrocarbon product having a ratio of even-numbered hydrocarbons toodd-numbered hydrocarbons of at least 2:1, preferably at least 3:1, morepreferably at least 5:1 and most preferably at least 10:1. Preferably,the catalyst comprises nickel, cobalt and molybdenum. When the renewableresource comprises over 50% C₁₈-based components, such as triglycerides,the process generally favors hydrodeoxygenation rather thandecarbonylation and/or decarboxylation.

Surprisingly, use of non-precious metals such as nickel, cobalt,molybdenum, or combinations thereof in the hydrotreating process of thisinvention produces yields of hydrotreated product equivalent or betteryields produced using the more expensive, precious metal catalysts suchas disclosed in U.S. Patent Publication 2006/0207166. Yields of 90% orbetter by volume can be achieved with the desired ratio of C₁₇ to C₁₈ ofbetween 0.05 to 0.95 (based on the desire to reject oxygen either as CO₂or H₂O) and the desired ratio of branched (isomerized) to the normal(straight) hydrocarbon can be varied from 0.05 to 0.95% based on thedesire for low pour point (more branched) or the desire for increasedcetane number (more straight chain).

The present invention may be performed in any suitable type of reactor.Suitable reactors include a fixed bed reactor and a slurry reactor. Afixed bed reactor has an advantage of easy separation of the reactantsand products from the catalyst. Fixed bed reactors include plug flow andtrickle bed reactors. Fixed bed reactors can be of the type adiabatic,multi-tubular, continuous recirculating packed bed reactor. Slurryreactors include batch, a continuously stirred tank reactor, and abubble column reactor. In the slurry reactors, the catalyst may beremoved from the reaction mixture by filtration or centrifugal action.Preferably, the process of this invention is a continuous process andthe reactor is a fixed bed or continuously stirred tank reactor. Morepreferably, the process is a continuous process and the reactor is afixed bed reactor.

Preferably, the process is a continuous process in a fixed bed or slurryreactor and the catalyst is in the form of particles, preferably shapedparticles. By “shaped particle” it is meant the catalyst is in the formof an extrudate. Extrudates include cylinders, pellets, or spheres.Cylinder shapes may have hollow interiors with one or more reinforcingribs. Trilobe, cloverleaf, rectangular- and triangular-shaped tubes,cross, and “C”-shaped catalysts can be used.

Preferably the shaped catalyst particle is about 0.01 to about 0.5 inch(about 0.25 to about 13 mm) in diameter when a packed bed reactor isused. More preferably, the catalyst particle is about 1/32 to about ¼inch (about 0.79 to about 6.4 mm) in diameter.

A wide range of suitable catalyst concentrations may be used. The amountof catalyst per reactor is generally dependent on the reactor type. Fora fixed bed reactor, the volume of catalyst per reactor will be high,while in a slurry, the volume will be lower. Typically, in a slurryreactor, the catalyst will make up 0.1 to about 30 wt % of the reactorcontents. Preferably, the catalyst is 1 to 15 wt % of the reactorcontents.

For a fixed bed reactor, the weight hourly space velocity will typicallyfall in the range of 0.05 to 100 hr⁻¹, preferably, 0.1 to 10 hr⁻¹, morepreferably 1.0 to 5.0 hr⁻¹.

In one embodiment of the process of the present invention, the feed iscontacted with hydrogen to form a liquid feed/hydrogen mixture inadvance of contacting the liquid feed/hydrogen mixture with thecatalyst. Optionally, a solvent or diluent, having a relatively highsolubility for hydrogen so that substantially all the hydrogen is insolution, can be added to the feed and hydrogen in advance of contactingwith the catalyst to form a liquid feed/solvent or liquid feed/diluentmixture. The liquid feed/solvent or liquid feed/diluent mixture is thencontacted with hydrogen to form a liquid feed/solvent/hydrogen or liquidfeed/diluent/hydrogen mixture. The mixture containing hydrogen is thencontacted with the catalyst.

In a preferred process, the liquid feed/solvent/hydrogen or liquidfeed/diluent/hydrogen mixture is contacted with catalyst in a packed bedreactor, such as plug flow, tubular or other fixed bed reactor for feedand hydrogen to react. It should be understood that the packed bedreactor may be a single packed bed or multiple beds in series or inparallel or in a combination thereof as discussed hereinabove.

The liquid feed/solvent/hydrogen or liquid feed/diluent/hydrogen mixturecan be a substantially hydrogen-gas-free liquid feed stream. The feedstream can be produced by contacting liquid feed with hydrogen andsolvent or diluent to produce a hydrogen-saturated liquid feed.Alternatively or in addition, after contacting liquid feed with hydrogenand solvent or diluent, hydrogen gas can be removed from the feedstream, for example, by known gas/liquid separation methods in adisengagement step. Processes for producing hydrogen-gas-free liquidfeed streams are known, such as those disclosed in U.S. Pat. Nos.6,123,835; 6,428,686; 6,881,326 and 7,291,257.

The percentage of hydrogen soluble in the solvent/diluent is greaterthan the percentage of hydrogen soluble in the liquid feed reactant. Inthis embodiment, preferably all of the hydrogen required for reaction ismade available in solution upstream of the fixed bed reactor, thuseliminating the need to circulate hydrogen gas within the reactor.

The reaction of liquid feed/solvent/hydrogen or liquidfeed/diluent/hydrogen mixture with catalyst is highly exothermic and asa result a great deal of heat is generated in the reactor. Thetemperature of the reactor can be controlled by using a recycle stream.A portion of the paraffin (hydrocarbon) product, (reactor effluent) canbe recycled back to the front of the reactor as a recycle stream andblended with fresh feed and hydrogen for use as solvent or diluent.

The process can be a multi-stage process using a series of two or morereactors in series and fresh hydrogen can be added at the inlet of eachreactor. The recycle stream absorbs some of the heat and reduces thetemperature rise through the reactor. The reactor temperature can becontrolled by controlling the fresh feed temperature and the amount ofrecycle. In addition, because the recycle stream comprises reactedcomponents, the recycle stream also serves as an inert diluent.

The type and amount of diluent added, as well as the reactor conditionscan be set so that substantially all of the hydrogen required in thehydrotreating reactions is available in solution. The solvent or diluentis preferably a portion of the reactor effluent used as a recyclestream, but can alternatively be selected from the group consisting oflight hydrocarbons, light distillates, naphtha, diesel, or the like.Examples include propane, butanes, and/or pentanes. The percentage ofhydrogen in the solvent or diluent is greater than the percentage ofhydrogen in the feed, thus, in this embodiment, all of the hydrogenrequired for reaction is made available in solution upstream of thereactor and eliminating the need to re-circulate hydrogen gas co-elutingwith the effluent or product stream.

Green Diesel

The process of this invention may be used to produce green diesel. Greendiesel produced in the process of this invention has the desiredproperties for use as diesel or for blending with petro diesel. Thesubstantially linear product has a high cetane number, which is neededto maintain power for diesel engines to run efficiently. One can use theproduct as fuel alone, or to blend in lower cetane products, such aslight cycle oil, oil sands or kerosene. (Light cycle oils can not beused as a diesel fuel without the used of cetane enhancing additives.)

The green diesel produced in this invention raises the cetane numberwithout negatively impacting the density. Cetane numbers can becontrolled by the selection of the specific catalyst and the processconditions. Cetane numbers are desired to be in the range of 50 to 100,more preferably 70 to 100. The branching of some of the chains and thecracking into smaller chains lowers the cloud point temperatures thatwould allow its usage in cold weather applications down to −40° C., whenblended in cold climate petro diesel. The degree of branching isdependent on the temperature of the application and can be controlled bythe selection of the properties of the zeolites used in the process andthe type and the combination of the metals. Green diesel produced bythis process also exhibit the desired lubricity (400 to 650 microns),viscosity (3 to 5 cSt at 40° C.), and density (750 to 800 kg m³ at 25°C.) suitable for today's diesel engines.

The present invention provides a process for more economical productionand implementation of green diesel with little or no impact on currentrefining production facilities or current diesel engines.

EXAMPLES Analytical Methods

Samples were dissolved in chloroform and analyzed by GC/MS for peakidentification and by GC/FID quantify the individual components unlessotherwise stated. For these examples, it is presumed the FID peak area %is equally comparable to weight percent due to the similarity ofcomponents (all hydrocarbons, linear and branched). Peaks less than 0.1%were neither identified nor quantified hence the totals are 90% to 95%range. In Table 1, ethanol and a few other minor impurities arecontaminants in the chloroform solvent. Most of the yield loss in allthe examples, but especially in Examples 5 through 8, is due toevaporation of low boilers upon depressuring the reaction tubes and dueto some residual product remaining in the tubes after the reaction wascompleted and the vessel was emptied.

Catalyst Sulfiding Procedure

A reactor consisting of ¾″ (19 mm) OD 316L Stainless Steel tubing 14″(36 cm) long was used for sulfiding catalyst. The reactor was packedwith alternating layers of 1 mm glass beads and PYREX wool at both ends,except in the middle, where catalyst (10 to 30 g) was packed. Thereactor had 3 thermocouples measuring the gas inlet, gas outlet, andcatalyst bed temperatures. The reactor was placed in the vertical tubefurnace and the gas inlet and the gas outlet connections wereestablished. The catalyst was allowed to dry overnight at 130° C. with a200 sccm of nitrogen flow. After drying the catalyst, the oventemperature was increased at a rate of 0.5-1.0° C./minute and 20 sccmhydrogen sulfide (5% mixture in hydrogen) was added to the 200 sccm N₂flow. Once the temperature reached 190° C., the nitrogen flow wasreduced to 100 sccm and the hydrogen sulfide flow was increased to 30sccm. The temperature was held at 240° C. After 2 hours, the temperaturewas reduced slowly. Once the temperature was below 125° C., the hydrogensulfide flow was stopped but the nitrogen flow was maintained at 100sccm until reactor reached room temperature (approx 25° C.). The reactorwas removed from the furnace and was unloaded in a nitrogen purge box.

Catalyst Reduction Procedure

Similar equipment and set-up used for sulfiding catalyst was used forreducing the catalyst. The catalyst was dried overnight at 130° C. undera 200 sccm nitrogen flow. The reactor had 3 thermocouples measuring thegas inlet, gas outlet, and catalyst bed temperatures. The reactor wasplaced in the vertical tube furnace and the gas inlet and the gas outletconnections were established. The catalyst was allowed to dry overnightat 130° C. with a 200 sccm of nitrogen flow. After drying the catalyst,the oven temperature was increased at a rate of 0.5-1.0° C./minute and20 sccm hydrogen gas (99.0% purity) was added to the 200 sccm N₂ flow.Once the temperature reached 190° C., the nitrogen flow was reduced to100 sccm and the hydrogen flow was increased to 30 sccm. The temperaturewas held at 240° C. unless noted otherwise. For Comparative Examples Band D the temperature was increased to 250° C. and for ComparativeExample C, the temperature was increased to 350° C. After 2 hours, thetemperature was reduced slowly. Once the temperature was below 125° C.,the hydrogen flow was stopped but the nitrogen flow was maintained at100 sccm until reactor reached room temperature (approx 25° C.). Thereactor was removed from the furnace and was unloaded in a nitrogenpurge box.

Example 1

Soybean oil (100 g, available from Sigma-Aldrich Co., St. Louis, Mo.)and a reduced Ni/NiO/MgO/SiO₂/graphite catalyst (Pricat Ni 55/5 P, >30wt % Ni metal, >25 wt % NiO, 5 g, available from Johnson Matthey, WestDeptford, N.J.) were placed in a 400 cc agitated pressure reactor.Soybean oil comprises triglycerides with the following distributionchain lengths: C₁₂=5%, C₁₄=5%, C₁₆=10%, C₁₈=3%, C_(18:1)=20%,C_(18:2)=50%, C_(18:3)=7%. C_(18:1) refers to an 18 carbon chain with 1unsaturated bonds, C_(18:2) refers to an 18 carbon chain with 2unsaturated bonds, and C_(18:3) refers to an 18 carbon chain with 3unsaturated bonds. The autoclave headspace was purged first withnitrogen 10 times by pressurizing/depressurizing between 90 and 0 psig(722 and 101 kPa), then with industrial grade hydrogen (high pressure99% purity, available from GTS Inc., Morrisville, Pa., USA) 5 times, andfinally pressurized to 500 psig (3550 kPa) with hydrogen. The autoclaveand its contents were heated to 250° C. with agitation. The hydrogenpressure was increased to 2000 psig (13,900 kPa), and maintained therefor 5 hours. The headspace was filled with fresh hydrogen to 2000 psig(13,900 kPa) if the pressure dropped below 1500 psig (10,400 kPa). Thetemperature was maintained at 250±10° C.

The autoclave contents were then cooled to below 50° C., the headspacewas vented, and the contents (103 g) were discharged to a glass bottle.IR and ¹H NMR analysis showed the presence of residual mono-, di- andtriglycerides. A small sample was then transesterified using methanoland base catalyst and the resulting sample was analyzed using GC-MS(peak identification) and GC-FID (species quantification) to show thatthe sample was equally divided among the methyl esters of the fattyacids (unreacted glycerides), alpha-olefins and normal paraffins(hydrocarbons). Thus, the nickel catalyst hydrodeoxygenated the soybeanoil.

Example 2

Soybean oil (50 g) and the catalyst used in Example 1 were placed in a400 cc agitated pressure reactor. The reaction was run at 300° C. andthe catalyst contained USY zeolite powder (0.125 g, type EZ-190,available from Engelhard (now part of BASF), Si/Al=3.05) physicallymixed in. The reaction contents were weighed (51 g). The sample was basetransesterified. IR showed the sample to be pure hydrocarbon with atrace of ester. A proton NMR analysis showed that the ester impurity wasminute (<100 ppm). A GC-FID analysis gave the following linear paraffin(hydrocarbon) product distribution by weight: C₁₈₊=1%, C₁₈=2%, C₁₇=78%,C₁₆=3%, C₁₅=11%, C₁₄=1%, C¹⁴⁻=4%. Some branching (<0.5 wt %isoheptadecane, “iso-C₁₇” was observed.

Example 3

The process of Example 2 was repeated using the same equipment, pressureand temperature conditions, and the reactants except for the catalystand no zeolite was added. The catalyst used was reduced nickel powder(45 wt % Ni metal, 24 wt % NiO catalyst on zirconia and kieselguhr(E-473P, 2.5 g, available from BASF Catalysts, Houston, Tex., USA). Thereaction products were weighed (51 g). An IR spectrum showed no ester inthe sample. A GC-FID analysis gave the following linear paraffin(hydrocarbon) distribution by weight: C₁₈₊=1%, C₁₈=2%, C₁₇=84%, C₁₆=1%,C₁₅=11%, C¹⁴⁻=1%. Branching was not observed.

Example 4

Example 2 was repeated using the same equipment, pressure andtemperature conditions. The reactants were the same as in Example 2, butdifferent amounts were used (100 g soybean oil, 5 g Ni 55/5 P catalyst,and 0.5 g of USY zeolite powder, type EZ-190). A GC-FID analysis gavethe following product distribution by weight: n-C₁₈ acid (octadecanoicacid)=56%, n-C₁₈=9%, n-C₁₇=29%, iso-C₁₇=1%, n-C16=1%, n-C15=4%. TheExample shows that the addition of zeolite catalyst results inisomerized product, with low, 10% zeolite present.

Examples 5 through 8

Examples 5 through 8 illustrate hydrocracking and hydroisomerization ina single step. Results from Examples 5 through 8 are summarized in Table1 below. The Examples used a feed comprised of a mixture ofhydrocarbons: 3% n-C₁₉, 91% n-C₁₈, and 6% n-C₁₇) prepared byhydrodeoxygenating canola oil in a continuous flow reactor using acommercial nickel/molybdenum on alumina catalyst at a temperature 325°C. and pressure 1500-2000 psig (10,400-13,900 kPa), followed bydistilling the product to obtain a predominantly n-C₁₈ cut.

This feed mixture (100 g) was reacted with reducedNi/NiO/MgO/SiO₂/graphite catalyst (Pricat Ni 55/5 P catalyst) (2.5 g)individually mixed with 4 different zeolite powders (2.5 g), specifiedin Table 1. These Examples were conducted at 300° C. and 1500 to 2000psig (13,900 kPa) under hydrogen in 400-cc pressure tubes for 5 hoursunder constant shaking. H₂ uptake was less than those in Examples 1through 4.

TABLE 1 Zeolites and weight percent of products for Examples 5 through 8Example No. 5 6 7 8 Zeolite Mordenite Beta LZ-Y-84 USY Product Amount, gProduct I.D. by GC/MS 87 82 42 48 Isobutane 4.4 6.1 9.8 8.9 Isopentane12.6 14.1 35.2 31.8 Isohexane 12.9 14.5 37.1 34.8 n-Heptane 5.6 5.7 6.48.0 2-Methylheptane 4.9 3.4 — 1.3 3-Methylheptane 3.5 2.9 — 1.13-Methyloctane 4.6 3.3 — 4.0 Ethanol 11.4 7.3 3.5 — 2-Methylnonane 3.72.4 — — p-Xylene — — 0.1 0.1 Undecane — — — — o-Xylene — — — 0.1Dodecane — — — — 1,2,4-Trimethylbenzene — — — 0.1 C3-Benzene — — — 0.2Tridecane — — 0.2 — Tetradecane — — — — Tetradecane C14 — — — —Pentadecane C15 — — — — Hexadecane C16 8.2 3.8 — — Heptadecane C₁₇ 3.53.7 — — 4-Ethyltetradecane 2.1 1.0 — — 4-Methylhexadecane 0.7 0.3 — —2-Methylheptadecane 0.4 0.2 — — 4-Methylheptadecane 1.6 0.9 — —Octadecane C₁₈ 14.2 25.7 — — Nonadecane C₁₉ 0.3 0.6 — — Percent Totals94.6 95.9 92.3 90.4

As can be seen from table 1, linear hydrocarbons similar to thoseproduced in Examples 1 through 3, undergo hydroisomerization andhydrocracking using catalysts comprising nickel supported on aluminacombined with a zeolite.

Example 9

n-Octadecane feed (100 g, same as used in Examples 5 through 8) andreduced Ni/NiO/MgO/SiO₂/graphite catalyst (Pricat Ni 55/5 P catalyst,2.5 g) were placed in a 400 cc agitated pressure reactor. The autoclaveheadspace was purged first with nitrogen 10 times bypressurizing/depressurizing between 90 and 0 psig (722 and 101 kPa),then with industrial grade hydrogen (high pressure, 99% purity,available from GTS Inc., Morrisville, Pa.) 5 times, and finallypressurized to 500 psig (3550 kPa) with hydrogen. The autoclave and itscontents were heated to 325° C. with agitation. The hydrogen pressurewas increased to 2000 psig (13,900 kPa), and maintained there for 5 hrs.The headspace was filled with fresh hydrogen to 2000 psig (13,900 kPa)if the pressure dropped below 1500 psig (10,400 kPa). The temperaturewas maintained at 325±10° C.

The autoclave contents were then cooled to below 50° C., the headspacewas vented, and the contents (˜100 g, including the catalyst) weredischarged to a glass bottle. A GC-FID analysis gave the followinglinear paraffin (hydrocarbon) product distribution by weight: C₁₉=3%,C₁₈=84%, C₁₇=13%. Note that the catalyst seemed to have converted somen-C₁₈ to n-C₁₇ through hydrocracking, but did not cause anyisomerization.

Example 10

Soybean oil from Sigma-Aldrich, 50 g, and alumina-supported pre-sulfidedcobalt/nickel/molybdenum tri-metallic hydrotreating catalyst (5 g, CRIDC2318, commercially available Criterion Catalysts and Technologies,Houston, Tex.) were placed in a 210 cc agitated pressure reactor. Thevessel was leaked check with nitrogen. The headspace of the reactor waspurged with nitrogen 10 times by pressurizing to 90 psig (722 kPa) anddepressurizing to 0 psig (101 kPa). The reactor was then purged withhigh purity hydrogen (99.9% min., commercially available from AirProducts, Allentown, Pa.) five times, and pressurized to 1000 psig (7000kPa) with hydrogen. The reactor and its contents were agitated andheated to 325° C. (617° F.). The hydrogen pressure was increased to 2000psig (13,900 kPa), and maintained there for 5 hours. The headspace wasfilled with fresh hydrogen to 1500-1700 psig (11,800 kPa) if thepressure dropped below 1000 psig (7000 kPa).

The reactor contents were then cooled to below 50° C. (122° F.), theheadspace was vented, and the contents were discharged to a glassbottle. The contents were weighed (51 g). IR and ¹H NMR analysis showedno evidence of mono-, di- and triglycerides. The sample was thenanalyzed using GC-MS (peak identification) and GC-FID (speciesquantification) to show that the sample was converted to the followinglinear paraffins (hydrocarbons) by weight: C₁₈ ₊ =1%, C₁₈=81%, C₁₇=6%,C₁₆=11.5%, C₁₅=0.5%. The C₁₈:C₁₇ ratio is greater than 13:1 and C₁₆:C₁₅ratio is greater than 20:1.

Example 11

Example 10 was repeated using the same equipment, reaction conditions,procedures and the reactant (50 g), except for different hydrotreatingcatalyst (5 g, alumina-supported pre-sulfided nickel/molybdenumbi-metallic hydrotreating catalyst, CRI DN-3330, commercially availableCriterion Catalysts and Technologies, Houston, Tex.). A GC-FID analysisof the product gave the following linear paraffin (hydrocarbon) productdistribution by weight: C₁₈ ₊ =1%, C₁₈=73%, C₁₇=12%, C₁₆=10%, C₁₅=1.5%.There was also 2.5% of n-octadecanoic acid in the product. The C₁₈:C₁₇ratio is over 6 and C₁₆:C₁₅ ratio is over 7.

Example 12

Catalyst (5 g, CRI DC2318), temperature and pressure conditions wererepeated from Example 10 except crude soybean oil (50 g, obtained fromPerdue Farms, Salisbury, Md.) was used. The reaction products wereanalyzed by GC-FID to obtain the following linear paraffin (hydrocarbon)distribution by weight: C₁₈₊=0.5%, C₁₈=80%, C₁₇=7%, C₁₆=11.6%, C₁₅=0.9%.The C₁₈:C₁₇ ratio is greater than 11:1 and C₁₆:C₁₅ ratio is greater than12:1.

Example 13

Catalyst (5 g, CRI DC2318), temperature and pressure conditions wererepeated from Example 10 using the same equipment, pressure, andtemperature, except for a refined, bleached, and deodorized soybean oilsample (50 g, obtained from Perdue Farms, Salisbury, Md.) was used. Thereaction products were analyzed by GC-FID to obtain the following linearparaffin (hydrocarbon) distribution by weight: C₁₈₊=0.5%, C₁₈=79%,C₁₇=8%, C₁₆=11.3%, C₁₅=1.2%. The C₁₈:C₁₇ ratio is almost 10, and C₁₆:C₁₅ratio is greater than 9.

Example 14

The process of Example 13 was repeated using the same equipment,pressure, temperature, and catalyst (5 g), except refined coconut oil,(50 g, obtained from Spectrum Chemicals of Gardena, Calif.) was used.The reaction products were analyzed by GC-FID to obtain the followinglinear paraffin (hydrocarbon) distribution by weight: C₁₈=9%, C₁₇=1%,C₁₆=9%, C₁₅=1%, C₁₄=17.5, C₁₃=2, C₁₂=43.5%, C₁₁=4, C₁₀=5.5, C₉=0.5,C₈=6.5%, C₇=0.5%. The C₁₈:C₁₇ ratio is approximately 9, C₁₆:C₁₅ ratio is9, C₁₄:C₁₃ ratio is approximately 9, C₁₂:C₁₁ ratio is approximately 11,C₁₀:C₉ ratio is 11, and C₈:C₇ ratio is 13.

Example 15

The process of Example 13 was repeated using the same equipment,pressure, temperature, and catalyst (5 g), except palm oil (50 g,manufactured by T.I. International Ghana Ltd. of Accra, Ghana) was used.The reaction products were analyzed by GC-FID to obtain the followinglinear paraffin (hydrocarbon) distribution by weight: C₁₈₊=0.5%,C₁₈=46.5%, C₁₇=5%, C₁₆=43%, C₁₅=4%, C₁₄=1%. The C₁₈:C₁₇ ratio is greaterthan 9, and C₁₆:C₁₅ ratio is greater than 10.

Example 16

The process of Example 13 was repeated using the same equipment,pressure, temperature, and catalyst (5 g), except chicken fat (50 g,obtained from Perdue Farms of Salisbury, Md.) was used. The reactionproducts were analyzed by GC-FID to obtain the following linear paraffin(hydrocarbon) distribution by weight: C₁₈₊=1%, C₁₈=60%, C₁₇=7%, C₁₆=28%,C₁₅=3%, C₁₄=1%. The C₁₈:C₁₇ ratio is approximately 9, and C₁₆:C₁₅ ratiois greater than 9.

Example 17 D100824-67 Stearic Acid , CRI Catalyst DC-2318

The process of Example 13 was repeated using the same equipment,pressure, temperature, and catalyst (5 g), except stearic acid (50 g,from VWR, West Chester, Pa.) was used. The reaction products wereanalyzed by GC-FID to obtain the following linear paraffin (hydrocarbon)distribution by weight: C₁₈₊=0.7%, C₁₈=90.4%, C₁₇=8.4%, C₁₆=0.5%. TheC₁₈:C₁₇ ratio is approximately 10.8.

Example 18

The process of Example 13 was repeated using the same equipment,pressure, temperature, and catalyst (5 g), except a 50:50 chicken fat tosoybean oil mixture (50 g, mixed in-house with chicken fat and soybeanoil obtained from Perdue Farms of Salisbury, Md.) was used. The reactionproducts were analyzed by GC-FID to obtain the following linear paraffin(hydrocarbon) distribution by weight: C₁₈₊=2.1%, C₁₈=69.6%, C₁₇=7.2%,C₁₆=18.5%, C₁₅=1.9%, C₁₄=0.5%, C₁₃=0.1% C₁₂=0.1%. The C₁₈:C₁₇ ratio isapproximately 9.7, and C₁₆:C₁₅ ratio is greater than 2.5.

Comparative Example A

The process of Example 18 was repeated using the same equipment,pressure, except the temperature was 400° C., the catalyst wasalumina-supported non-sulfided nickel/molybdenum bimetallichydrotreating catalyst, (5 g, AT 535, from Grace-Davidson, Columbia,Md.), and crude soybean oil (50 g, obtained from Perdue Farms,Salisbury, Md.) was used. The catalyst was reduced, as describedhereinabove. No reaction occurred.

Comparative Example B

The process of Example 18 was repeated using the same equipment,pressure, except the temperature was 500° C., Grace-Davidson AT 535catalyst (5 g), and crude soybean oil from Perdue Farms (50 g) was used.The catalyst was reduced, as described herein above. No reactionoccurred.

Example 19

The process of Example 18 was repeated using the same equipment,pressure, and temperature except the catalyst was Grace-Davidson AT 535catalyst (5 g) and crude soybean oil from Perdue Farms (50 g) was used.The catalyst was sulfided, as described hereinabove. The reactionproducts were analyzed by GC-FID to obtain the following linear paraffindistribution by weight: C₁₈₊=2.0%, C₁₈=78.0%, C₁₇=8%, C₁₆=10%, C₁₅=1.1%,C₁₄=0.2%, C₁₃=0.2%, C₁₂=0.1%, C₁₁=0.1%, C₇₋₁₀=0.3%. The C₁₈:C₁₇ ratio isapproximately 9.75, and C₁₆:C₁₅ ratio is greater than 9.

Comparative Example D

The process of Example 18 was repeated using the same equipment,pressure, and temperature, except the catalyst was alumina-supportednon-sulfided nickel/cobalt/molybdenum tri-metallic hydrotreatingcatalyst, (5 g, AT 592, from Grace-Davidson, Columbia, Md.), and crudesoybean oil (50 g, obtained from Perdue Farms, Salisbury, Md.) was used.The catalyst was reduced, as described hereinabove. No reactionoccurred.

Example 20

The process of Example 19 was repeated using the same equipment,pressure, and temperature except Grace-Davidson AT 792 catalyst (5 g,)was used. The catalyst was sulfided, as described hereinabove. Thereaction products were analyzed by GC-FID to obtain the following linearparaffin distribution by weight: C₁₈₊=1.8%, C₁₈=82.4%, C₁₇=3.3%,C₁₆=11.2%, C₁₅=0.5%, C₁₄=0.2%, C₁₃=0.1%, C₁₂=0.1%, C₁₁=0.1%, C₇₋₁₀=0.3%.The C₁₈:C₁₇ ratio is approximately 25, and C₁₆:C₁₅ ratio is greater than22.

As can be seen from Examples 10-20 and Comparative Examples A-D, thecobalt/nickel/molybdenum and nickel/molybdenum catalysts produced higherratio of even-numbered hydrocarbons to odd-numbered hydrocarbons.

1. A process for hydrodeoxygenation of a renewable resource whichcomprises (a) providing a feed which is a renewable resource; (b)contacting the feed with a catalyst in the presence of hydrogen at atemperature of 250 to 425° C. and a pressure of 500 to 2500 psig (3,450to 17,250 kPa), wherein the catalyst comprises an oxide, molybdenum andone or more active metals selected from the group consisting of nickel,cobalt, and mixtures thereof and the catalyst is sulfided form, toproduce a hydrocarbon product having a ratio of even-numberedhydrocarbons to odd-numbered hydrocarbons of at least 2:1.
 2. Theprocess of claim 1 wherein the feed is an oil derived from plants and/oranimals and comprises one or more triglyceride or one or more free fattyacid.
 3. The process of claim 2 wherein the triglyceride is derived froma plant selected from the group consisting of pine, rape seed,sunflower, palm, jathropa, seashore mallow and combinations of two ormore thereof.
 4. The process of claim 1 wherein the feed is a vegetableoil selected from the group consisting of canola oil, palm oil, coconutoil, palm kernel oil, sunflower oil, soybean oil, crude tall oil, andcombinations of two or more thereof.
 5. The process of claim 1 whereinthe feed is poultry fat, yellow grease, or tallow.
 6. The process ofclaim 1 wherein the concentration of metal in the catalyst is 0.1 to 90percent by weight, based on the total weight of the catalyst.
 7. Theprocess of claim 6 wherein the concentration of metal in the catalyst is0.5 to 60 weight percent.
 8. The process of claim 1 wherein the oxide isselected from the group consisting of alumina, silica, titania,zirconia, kieselguhr, silica-alumina and combinations.
 9. The process ofclaim 1 wherein the temperature is 300 to 400° C.
 10. The process ofclaim 9 wherein the temperature is 325 to 375° C.
 11. The process ofclaim 1 wherein the pressure is 1000 to 2000 psig (7000 to 13,900 kPa).12. The process of claim 1 wherein the ratio of even-numberedhydrocarbons to odd-numbered hydrocarbons of at least 3:1.
 13. Theprocess of claim 1 wherein the ratio of even-numbered hydrocarbons toodd-numbered hydrocarbons of at least 5:1.
 14. The process of claim 1wherein the ratio of even-numbered hydrocarbons to odd-numberedhydrocarbons of at least 10:1.
 15. The process of claim 1 wherein theprocess is a continuous process and the reactor is a fixed bed reactoror a continuously stirred tank reactor.
 16. The process of claim 15wherein the reactor is a fixed bed reactor.
 17. The process of claim 15wherein the feed is contacted with a solvent or diluent and hydrogen toprovide a liquid feed/solvent/hydrogen or liquid feed/diluent/hydrogenmixture in advance of contacting the feed with the catalyst.
 18. Theprocess of claim 17 wherein the liquid feed/solvent/hydrogen or liquidfeed/diluent/hydrogen mixture is a substantially hydrogen-gas-freeliquid feed stream.
 19. The process of claim 17 wherein a portion of theproduct is recycled back to the reactor as a recycle stream and blendedwith fresh feed and hydrogen for use as solvent or diluent.